Catalyst return apparatus and process for reacting a feedstock

ABSTRACT

A catalyst return apparatus is disclosed as well as a riser reactor system comprising the conduit apparatus and a riser reactor, the conduit apparatus comprising a catalyst return conduit and at least two flow control devices in series, each flow control device arranged to control the flow of fluid through the conduit, wherein the length of the catalyst return conduit is more than 20 m. A process for reacting a feedstock in a riser reactor system comprising a riser reactor, the catalyst return apparatus and, and a stage vessel, the process comprising: holding a fluid comprising the catalyst in the at least one stage vessel for a residence time of at least 10 seconds.

PRIORITY CLAIM

This is a divisional application of U.S. patent application Ser. No.13/141,002, filed Aug. 24, 2011, which claims priority fromPCT/EP2009/067686, filed 21 Dec. 2009, which claims the benefit ofpriority from European Application 08172543.4, filed 22 Dec. 2008, eachof which is incorporated herein by reference.

BACKGROUND

This invention relates to a riser reactor system, a catalyst returnconduit apparatus and a process for reacting a feedstock within theriser reactor system, especially an oxygenate feedstock to produceolefins.

Processes for the preparation of olefins from oxygenates are known inthe art. Of particular interest is often the production of lightolefins, in particular ethylene and/or propylene. The oxygenatefeedstock can for example comprise methanol and/or dimethylether, and aninteresting route includes their production from synthesis gas derivedfrom e.g. natural gas or via coal gasification.

For example, WO2007/135052 discloses a process wherein an alcohol and/orether containing oxygenate feedstock and an olefinic co-feed are reactedin the presence of a zeolite having one-dimensional 10-membered ringchannels to prepare an olefinic reaction mixture, and wherein part ofthe obtained olefinic reaction mixture is recycled as olefinic co-feed.With a methanol and/or dimethylether containing feedstock, and anolefinic co-feed comprising C4 and/or C5 olefins, an olefinic productrich in light olefins can be obtained.

A suitable reactor system for oxygenate-to-olefins reactions includes ariser reactor. Using a riser reactor, a continuous process can beemployed where the used catalyst is separated from the product and otherfluids in a separation zone, at least some of the catalyst isregenerated to remove some of the coke deposits, and catalyst isreintroduced into the riser reactor via a catalyst return conduit, alsoreferred to as a standpipe. It is important for smooth continuousoperation of a riser reactor system that the solids circulation operateswell.

US2003/0234209 teaches a method for controlling solids circulation in agas/solids reaction system. The method entails aerating solid particlesin a standpipe, wherein aeration fluid is injected into the standpipeaway from the internal wall.

In the design of riser reactor systems, e.g. for oxygenate-to-olefinsconversions, it can be desirable to use tall reactors, in order toprovide the right reactions conditions in terms of catalystconcentration, superficial velocity, flow regime and/or residence timefor a given cross-sectional area of the riser. For commerciallyinteresting throughputs and a desired height aspect ratio between heightand cross-sectional area, reactors can be desired to reach heights of 30m, 50 m, 70 m or even more.

US2004/0076554 addresses a particular problem encountered when designinga reactor system so tall, namely that heavy equipment at the top of sucha structure requires expensive support structures. US2004/0076554discloses a multiple riser reactor, wherein the effluent from risers islaterally fed into a common separation vessel, which is arranged betweenthe riser reactors and not on above them. From the lower end of theseparation vessel catalyst guided via a central downcomer to a catalystretention zone, and from the catalyst retention zone short standpipesfeed the catalyst back to the inlets of the multiple risers. In this waythe overall height of the structure is reduced.

Applicant has realized another problem in designing tall riser reactors,in particular when a design is used in which a catalyst retention zone,wherein catalyst is collected after separation and perhaps regeneration,is arranged at a very high position, e.g. at the top of a tall riserreactor. In this case the standpipe for feeding catalyst from thecatalyst retention zone to the lower end of the riser becomes very tallas well. Applicant has realized that it is problematic to reliablycontrol the flow of fluid in such a standpipe of more than 20 metersheight, since valves that are typically used, e.g. slide valves, cannotbe used at higher pressure differences than about 1 bar. Otherwise wearwould be too high and control unreliable.

There is a need for an improved catalyst return in riser reactors.

SUMMARY OF THE INVENTION

According to a first aspect of the invention, there is provided acatalyst return apparatus suitable for use in a riser reactor system,the apparatus comprising a catalyst return conduit and at least two flowcontrol devices in series, wherein the length of the catalyst returnconduit is more than 20 m. Each flow control device is arranged tocontrol the flow of fluid through the catalyst return conduit. Theexpression ‘catalyst return conduit’ is user herein to denote a flowpathfor catalyst from a catalyst retention zone such as a collector vesselto an inlet of a riser reactor. The conduit can be formed of varioussections and elements that are in fluid communication. The catalystreturn conduit serves to return at least the larger part of the catalystto the reactor. Therefore, preferably, the catalyst return conduit doesnot comprise or incorporate regeneration or heat exchange units, as suchunits may put limitations on the catalyst throughput. Reference, hereinto heat exchangers is to units that would change the temperature of thecatalyst to an extent beyond the temperature change, occurring while thecatalyst is passed through the catalyst return conduit withoutadditional heat exchange

DETAILED DESCRIPTION OF THE INVENTION

By providing at least two flow control devices, such as valves, inseries, the height can be broken down in several portions, wherein eachof the valves only has to handle a portion of the overall pressure drop.In this way the pressure drop per valve can e.g. be limited to about 1bar or below.

The flow control devices are typically valves, and will in thedescription be referred to as valves for simplicity. The valves arenormally spaced apart by at least 5 m and preferably at least 10 m.Preferably at least one flow control device is present at least every 20m along the catalyst return apparatus.

The apparatus may further comprise a means to reconstitute catalystflow. Such means to reconstitute catalyst flow can in particular be astage vessel. The stage vessel has suitably a width at least twice awidth of the conduit. Typically the or each stage vessel has a width ofat least three, preferably at least four, times a width of the conduit.In use the stage vessel allows the returning fluidised catalyst toseparate into a predominantly gas phase above a predominantly fluidisedsolid phase. From the fluidized solid phase, a well-defined flow ofcatalyst into a following section or portion of the flowpath along thecatalyst return apparatus can start. The predominantly gas phase canhave more than 70 vol % gas, preferably more than 80 vol %. It will beclear that depending on the throughput the separation between the twophases may not be very sharp and that there may be a gradient ofdecreasing density from bottom to top, but solids will settle at thebottom and form a fluidize bed there. The predominantly fluidised solidphase can have more than 40 vol % solid, preferably more than 50 vol %solid.

Normally the stage vessel comprises a gas inlet to further fluidise thecatalyst during use.

Preferably the catalyst return conduit comprises at least one stagevessel at least every 20 m.

Preferably at least one flow control device, such as a valve, isprovided for the or each stage vessel, typically within 1 m upstream ofthe or each stage vessel, to control the fluid within the conduit atthat point. Typically the stage vessel is located between two valves.

The conduit may comprise a first portion above the stage vessel and asecond portion below the stage vessel, and further portions betweenstage vessels if more than one stage vessels is provided. The first andsecond portions and/or other portions are typically not co-linear. Eachportion of stand-pipe (conduit) may be at least 4 m in length.

The invention also provides a riser reactor system comprising a catalystreturn apparatus as described herein and a riser reactor.

The apparatus typically comprises at least one stage vessel, asdescribed herein.

The riser reactor system may comprise a pressure equalisation means toequalise the gas pressure in the stage vessel(s) with the gas pressureat the top of the riser reactor, especially with a separation zoneprovided at the top of the riser reactor. For example a further conduitmay be provided to connect the gas phase of the stage vessel(s) with thegas phase at the top of the riser reactor.

The invention also provides a process for reacting a feedstock in ariser reactor system with a catalyst return conduit according to theinvention and at least one stage vessel, the process comprising holdinga fluidised catalyst in the at least one stage vessel for a residencetime of at least 10 seconds, preferably at least 20 seconds.

The residence time may be determined by calculations based on the inputand output rate. The residence time can be defined as the average amountof catalyst in the stage vessel divided by the average flow of catalyst(in terms of amount per unit of time). The amount can for example bemeasured as a mass. The process of holding fluid within the stage vesselis very typically a continuous process rather than a batch process.

In the simplest embodiment of the riser reactor system, the riserreactor has a constant cross-sectional area. As riser reactors normallyhave circular cross-sections, their cross-sectional area is proportionalto their diameter.

In a preferred embodiment, the riser reactor system comprises two ormore serially arranged riser reactor stages, wherein each riser reactorstage comprises a single riser reactor segment or a plurality ofparallel riser reactors.

The applicant has found that selectivity of an oxygenate conversionprocess towards desired olefins, in particular ethylene, can besignificantly improved using a serial riser reactor system, in whichoxygenate feedstock together with fluidised catalyst is stagewise addedto a plurality of the riser reactor stages, in particular to the firstand at least one additional riser reactor stage.

Preferably each stage comprises a single riser reactor segment.Preferably therefore the riser reactor system comprises a plurality ofriser reactor segments.

In a particular embodiment, riser stages may be arranged by stackingriser segments on top of each other such that they are co-linear and sofluidised catalyst may flow up the lower riser of the first stage andthen continue up the upper riser of a subsequent stage. Such a stackedarrangement normally leads to a substantial overall height of thereactor system, such as more than 20 m, more than 30 m, or more than 40m, and so increases the benefit of the catalyst return conduit inaccordance with the present invention.

Preferably in such an embodiment a further catalyst return conduit isprovided to connect to a further catalyst inlet of a consecutive stageabove the catalyst inlet of the first stage. This further catalyst inletis typically at least 5m from the bottom of the riser reactor system.

Between 2 and 10, preferably between 2 and 5, more preferably three orfour, riser reactor stages, such as riser segments can be provided.

Preferably the first of the two or more serially arranged riser reactorstages has a smaller total cross-sectional area than at least one of thesubsequent riser reactor stages. It is particularly beneficial if thetotal cross-sectional area of each subsequent riser reactor stage, afterthe first, is higher than that of the preceding riser reactor stage. Thetotal cross-sectional area is the sum of the cross-sectional areas ofall riser reactors in a particular stage. In preferred embodiments whenthere is only one riser reactor segment in a stage, its cross-sectiondefines the total cross-sectional area of that stage. Increasing thecross-sectional area can partly or fully compensate for the increase involumetric flow rate due to additional catalyst (and optionallyfeedstock), so that the flow velocity in the riser does not increasebeyond critical values impeding for example conversion, catalyststability and/or attrition.

The invention also provides a process for the preparation of an olefinicproduct within a riser reactor system as described herein, the processcomprising reacting an oxygenate feedstock in the presence of anoxygenate conversion catalyst under oxygenate-to-olefin conversionconditions in the riser reactor system, to obtain the olefinic product.

Typically in use the oxygenate feedstock is contacted with an oxygenateconversion catalyst to obtain a riser reactor effluent from each stageand at least part of the riser reactor effluent of a preceding riserreactor stage is fed into a subsequent riser reactor stage, andpreferably fluidised catalyst and optionally oxygenate is added to aplurality of the riser reactor stages.

The riser reactor effluent of a preceding riser reactor stage comprisesgaseous effluent and solid oxygenate conversion catalyst. Normally, atleast 50 wt % of the gaseous effluent is fed to the subsequent riserreactor stage, in particular at least 80%, more in particular at least90%. Further, it can be beneficial not to separate solids and gasesbetween subsequent riser reactor stages. So, normally also at least 50wt % of the solid oxygenate conversion catalyst is fed to the subsequentriser reactor stage, in particular at least 80%, more in particular atleast 90%. More in particular, substantially all riser reactor effluentfrom one riser reactor stage can be fed to the subsequent riser reactorstage.

The oxygenate feedstock suitably comprises oxygenate species having anoxygen-bonded methyl group, such as methanol or dimethylether.Preferably the oxygenate feedstock comprises at least 50 wt % ofmethanol and/or dimethylether, more preferably at least 80 wt %, mostpreferably at least 90 wt %.

The oxygenate feedstock can be obtained from a different or separatereactor, which converts methanol at least partially into dimethyletherand water. Water may be removed by e.g. distillation. In this way, wateris present in the process of converting oxygenate to olefins, which hasadvantages for the process design and lowers the severity ofhydrothermal conditions the catalyst is exposed to.

The oxygenate feedstock can comprise an amount of water, preferably lessthan 10 wt %, more preferably less than 5 wt %. Preferably the oxygenatefeedstock contains essentially no hydrocarbons other than oxygenates,i.e. less than 5 wt %, preferably less than 1 wt %.

In one embodiment, the oxygenate is obtained as a reaction product ofsynthesis gas. Synthesis gas can for example be generated from fossilfuels, such as from natural gas or oil, or from the gasification ofcoal. Suitable processes for this purpose are for example discussed inIndustrial Organic Chemistry, Klaus Weissermehl and Hans-Jürgen Arpe,3rd edition, Wiley, 1997, pages 13-28. This book also describes themanufacture of methanol from synthesis gas on pages 28-30.

In another embodiment the oxygenate is obtained from biomaterials, suchas through fermentation. For example by a process as described inDE-A-10043644.

In one embodiment, oxygenate is added to each of the riser reactorstages. Preferably the oxygenate added at the different stages isderived from a common oxygenate feedstock source. The common oxygenatefeedstock source can for example be a storage vessel, feed line, or adifferent or separate reactor. In this way oxygenate comprisingfeedstock of substantially the same composition is fed to and convertedin each of the riser reactor stages.

In one embodiment, the mass flow rate of oxygenate conversion catalystin each subsequent riser reactor stage, after the first, to whichoxygenate is added is higher than in the preceding riser reactor stage.Thus, the addition of fresh oxygenate feed is accommodated by additionaloxygenate conversion catalyst. In this way the weight hourly spacevelocity (WHSV), defined as the throughput of the weight of reactantsand reaction products per hour, and per weight of catalyst in thereactor, can be maintained above a selected minimum value, in order toachieve sufficient conversion.

Preferably the oxygenate feedstock is reacted to produce the olefinicproduct in the presence of an olefinic co-feed. By an olefiniccomposition or stream, such as an olefinic product, product fraction,fraction, effluent, reaction effluent or the like is understood acomposition or stream comprising one or more olefins, unlessspecifically indicated otherwise. Other species can be present as well.Apart from olefins, the olefinic co-feed may contain other hydrocarboncompounds, such as for example paraffinic compounds. Preferably theolefinic co-feed comprises an olefinic portion of more than 50 wt %,more preferably more than 60 wt %, still more preferably more than 70 wt%, which olefinic portion consists of olefin(s). The olefinic co-feedcan also consist essentially of olefin(s).

Any non-olefinic compounds in the olefinic co-feed are preferablyparaffinic compounds. Such paraffinic compounds are preferably presentin an amount in the range of from 0 to 50 wt %, more preferably in therange of from 0 to 40 wt %, still more preferably in the range of from 0to 30 wt %.

By an olefin is understood an organic compound containing at least twocarbon atoms connected by a double bond. The olefin can be amono-olefin, having one double bond, or a poly-olefin, having two ormore double bonds. Preferably olefins present in the olefinic co-feedare mono-olefins. C4 olefins, also referred to as butenes (1-butene,2-butene, iso-butene, and/or butadiene), in particular C4 mono-olefins,are preferred components in the olefinic co-feed.

Preferably the olefinic co-feed is at least partially obtained by arecycle stream formed by recycling a suitable fraction of the reactionproduct comprising C4 olefin. The skilled artisan knows how to obtainsuch fractions from the olefinic reaction effluent such as bydistillation.

In one embodiment at least 70 wt % of the olefinic co-feed, duringnormal operation, is formed by the recycle stream, preferably at least90 wt %, more preferably at least 99 wt %. Most preferably the olefinicco-feed is during normal operation formed by the recycle stream, so thatthe process converts oxygenate feedstock to predominantly light olefinswithout the need for an external olefins stream. During normal operationmeans for example in the course of a continuous operation of theprocess, for at least 70% of the time on stream. The olefinic co-feedmay need to be obtained from an external source, such as from acatalytic cracking unit or from a naphtha cracker, during start-up ofthe process, when the reaction effluent comprises no or insufficient C4+olefins.

A particularly preferred olefinic recycle stream is a C4 fractioncontaining C4 olefin(s), but which can also contain a significant amountof other C4 hydrocarbon species, in particular C4 paraffins, because itis difficult to economically separate C4 olefins and paraffins, such asby distillation.

In a preferred embodiment the olefinic co-feed and preferably also therecycle stream comprises C4 olefins and less than 10 wt % of C5+hydrocarbon species, more preferably at least 50 wt % of C4 olefins, andat least a total of 70 wt % of C4 hydrocarbon species.

The olefinic co-feed and preferably also the recycle stream, can inparticular contain at least a total of 90 wt % of C4 hydrocarbonspecies. In a preferred embodiment, the olefinic co-feed comprises lessthan 5 wt % of C5+ olefins, preferably less than 2 wt % of C5+ olefins,even more preferably less than 1 wt % of C5+ olefins, and likewise therecycle stream. In another preferred embodiment, the olefinic co-feed,comprises less than 5 wt % of C5+ hydrocarbon species, preferably lessthan 2 wt % of C5+ hydrocarbon species even more preferably less than 1wt % of C5+ hydrocarbon species, and likewise the recycle stream.

Thus in certain preferred embodiments, the olefinic portion of theolefinic co-feed, and of the recycle stream, comprises at least 90 wt %of C4 olefins, more preferably at least 99 wt %. Butenes as co-feed havebeen found to be particularly beneficial for high ethylene selectivity.Therefore one particularly suitable recycle stream consists essentially,i.e. for at least 99 wt %, of 1-butene, 2-butene (cis and trans),isobutene, n-butane, isobutane, butadiene.

In certain embodiments, the recycle stream can also comprise propylene.This may be preferred when a particularly high production of ethylene isdesired, so that part or all of the propylene, such as at least 5 wt %thereof, produced is recycled together with C4 olefins.

The preferred molar ratio of oxygenate in the oxygenate feedstock toolefin in the olefinic co-feed depends on the specific oxygenate usedand the number of reactive oxygen-bonded alkyl groups therein.Preferably the molar ratio of oxygenate to olefin in the total feed liesin the range of 10:1 to 1:10, more preferably in the range of 5:1 to 1:5and still more preferably in the range of 3:1 to 1:3.

In a preferred embodiment wherein the oxygenate comprises only oneoxygen-bonded methyl group, such as methanol, the molar ratio preferablylies in the range of from 5:1 to 1:5 and more preferably in the range of2.5:1 to 1:2.5.

In another preferred embodiment wherein the oxygenate comprises twooxygen-bonded methyl groups, such as for example dimethylether, themolar ratio preferably lies in the range of from 5:2 to 1:10 and morepreferably in the range of 2:1 to 1:4. Most preferably the molar ratioin such a case is in the range of 1.5:1 to 1:3.

The process to prepare an olefin is typically carried out in presence ofa molecular sieve having one-dimensional 10-membered ring channels.These are understood to be molecular sieves having only 10-membered ringchannels in one direction which are not intersected by other 8, 10 or12-membered ring channels from another direction.

Preferably, the molecular sieve is selected from the group of TON-type(for example zeolite ZSM-22), MTT-type (for example zeolite ZSM-23),STF-type (for example SSZ-35), SFF-type (for example SSZ-44), EUO-type(for example ZSM-50), and EU-2-type molecular sieves or mixturesthereof.

MTT-type catalysts are more particularly described in e.g. U.S. Pat. No.4,076,842. For purposes of the present invention, MTT is considered toinclude its isotypes, e.g., ZSM-23, EU-13, ISI-4 and KZ-1.

TON-type molecular sieves are more particularly described in e.g. U.S.Pat. No. 4,556,477. For purposes of the present invention, TON isconsidered to include its isotypes, e.g., ZSM-22, Theta-1, ISI-1, KZ-2and NU-10.

EU-2-type molecular sieves are more particularly described in e.g. U.S.Pat. No. 4,397,827. For purposes of the present invention, EU-2 isconsidered to include its isotypes, e.g., ZSM-48.

In a further preferred embodiment a molecular sieve of the MTT-type,such as ZSM-23, and/or a TON-type, such as ZSM-22 is used.

Molecular sieve and zeolite types are for example defined in Ch.Baerlocher and L. B. McCusker, Database of Zeolite Structures:http://www.iza-structure.org/databases/, which database was designed andimplemented on behalf of the Structure Commission of the InternationalZeolite Association (IZA-SC), and based on the data of the 4th editionof the Atlas of Zeolite Structure Types (W. M. Meier, D. H. Olson andCh. Baerlocher). The Atlas of Zeolite Framework Types, 5th revisededition 2001 and 6^(th) edition 2007 may also be consulted.

Preferably, molecular sieves in the hydrogen form are used in theoxygenate conversion catalyst, e.g., HZSM-22, HZSM-23, and HZSM-48,HZSM-5. Preferably at least 50% w/w, more preferably at least 90% w/w,still more preferably at least 95% w/w and most preferably 100% of thetotal amount of molecular sieve used is in the hydrogen form. When themolecular sieves are prepared in the presence of organic cations themolecular sieve may be activated by heating in an inert or oxidativeatmosphere to remove organic cations, for example, by heating at atemperature over 500° C. for 1 hour or more. The zeolite is typicallyobtained in the sodium or potassium form. The hydrogen form can then beobtained by an ion exchange procedure with ammonium salts followed byanother heat treatment, for example in an inert or oxidative atmosphereat a temperature over 300° C. The molecular sieves obtained afterion-exchange are also referred to as being in the ammonium form.

Preferably the molecular sieve having one-dimensional 10-membered ringchannels has a silica to alumina ratio (SAR) in the range of from 1 to500, preferably in the range of from 10 to 200. The SAR is defined asthe molar ratio of SiO₂/Al₂O₃ corresponding to the composition of themolecular sieve.

For ZSM-22, a SAR in the range of 40-150 is preferred, in particular inthe range of 70-120. Good performance in terms of activity andselectivity has been observed with a SAR of about 100.

For ZSM-23, an SAR in the range of 20-120 is preferred, in particular inthe range of 30-80. Good performance in terms of activity andselectivity has been observed with a SAR of about 50.

In a special embodiment the reaction is performed in the presence of amore-dimensional molecular sieve, such as ZSM-5. Suitably to this endthe oxygenate conversion catalyst comprises at least 1 wt %, based ontotal molecular sieve in the oxygenate conversion catalyst, of a furthermolecular sieve having more-dimensional channels, in particular at least5 wt %, more in particular at least 8 wt %.

The further molecular sieve having more-dimensional channels isunderstood to have intersecting channels in at least two directions. So,for example, the channel structure is formed of substantially parallelchannels in a first direction, and substantially parallel channels in asecond direction, wherein channels in the first and second directionsintersect. Intersections with a further channel type are also possible.Preferably the channels in at least one of the directions are10-membered ring channels. The further molecular sieve can be forexample a PER type zeolite which is a two-dimensional structure and has8- and 10-membered rings intersecting each other. Preferably however theintersecting channels in the further molecular sieve are each10-membered ring channels. Thus the further molecular sieve may be azeolite, or a SAPO-type (silicoaluminophosphate) molecular sieve. Morepreferably however the further molecular sieve is a zeolite. A preferredfurther molecular sieve is an MFI-type zeolite, in particular zeoliteZSM-5.

The presence of the further molecular sieve in the oxygenate conversioncatalyst was found to improve stability (slower deactivation duringextended runs) and hydrothermal stability compared to a catalyst withonly the one-dimensional molecular sieve and without themore-dimensional molecular sieve. Without wishing to be bound by aparticular hypothesis or theory, it is presently believed that this isdue to the possibility for converting larger molecules by the furthermolecular sieve having more-dimensional channels, that were produced bythe first molecular sieve having one-dimensional 10-membered ringchannels, and which would otherwise form coke. When the one-dimensionalaluminosiclicate and the more-dimensional molecular sieve are formulatedsuch that they are present in the same catalyst particle, such as in aspray-dried particle, this intimate mix was found to improve theselectivity towards ethylene and propylene, more in particular towardsethylene.

The weight ratio between the molecular sieve having one-dimensional10-membered ring channels, and the further molecular sieve havingmore-dimensional channels can be in the range of from 1:100 to 100:1,preferably 1:1 to 100:1, more preferably in the range of 9:1 to 2:1.

Preferably the further molecular sieve is an MFI-type molecular sieve,in particular zeolite ZSM-5, having a silica to alumina ratio (SAR) ofat least 60, more preferably at least 80, even more preferably at least100, yet more preferably at least 150. At higher SAR the percentage ofC4 saturates in the C4 totals produced is minimized In specialembodiments the oxygenate conversion catalyst can comprise less than 35wt % of the further molecular sieve, based on the total molecular sievein the oxygenate conversion catalyst, in particular less than 20 wt %,more in particular less than 18 wt %, still more in particular less than15 wt %.

In one embodiment the oxygenate conversion catalyst can comprise morethan 50 wt %, at least 65 wt %, based on total molecular sieve in theoxygenate conversion catalyst, of the molecular sieve havingone-dimensional 10-membered ring channels. The presence of a majority ofsuch molecular sieve strongly determines the predominant reactionpathway.

The molecular sieve can be used as such or in a formulation, such as ina mixture or combination with a so-called binder material and/or afiller material, and optionally also with an active matrix component.Other components can also be present in the formulation. If one or moremolecular sieves are used as such, in particular when no binder, filler,or active matrix material is used, the molecular sieve itself is/arereferred to as oxygenate conversion catalyst. In a formulation, themolecular sieve in combination with the other components of the mixturesuch as binder and/or filler material is/are referred to as oxygenateconversion catalyst.

It is desirable to provide a catalyst having good mechanical or crushstrength, because in an industrial environment the catalyst is oftensubjected to rough handling, which tends to break down the catalyst intopowder-like material. The latter causes problems in the processing.Preferably the molecular sieve is therefore incorporated in a bindermaterial. Examples of suitable materials in a formulation include activeand inactive materials and synthetic or naturally occurring zeolites aswell as inorganic materials such as clays, silica, alumina,silica-alumina, titania, zirconia and aluminosilicate. For presentpurposes, inert materials, such as silica, are preferred because theymay prevent unwanted side reactions which may take place in case a moreacidic material, such as alumina or silica-alumina is used.

In one embodiment the oxygenate added can be used for temperaturecontrol, and to this end the temperature of the oxygenate added to atleast one of the riser reactors of any one of the riser reactor stagesis set in dependence of a predetermined desired temperature in thisriser reactor. For example, depending on the temperature and mass flowrate of the effluent stream from the previous riser reactor, thetemperature and mass flow rate of additional catalyst, the temperatureof the oxygenate can be set, e.g. by heat exchange, so that near theinlet of the riser reactor a predetermined inlet temperature of themixture of the various feeds is realized.

In one embodiment, each gaseous effluent from one of the riser reactorshas an oxygenate concentration below 10 wt %, in particular below 5 wt%, preferably below 2 wt %, more preferably below 1 wt %, still morepreferably below 0.1 wt %. In this way, substantially full conversion ofoxygenate in each riser reactor is realized. This is particularlybeneficial at the last reactor effluent, as otherwise unreactedoxygenate has to be separated from the effluent in a work-up section.Separating e.g. unreacted methanol from water formed in the process isan undesirable and costly step in an industrial process.

The reaction to produce the olefins can be carried out over a wide rangeof temperatures and pressures. Suitably, however, the oxygenate feed andolefinic co-feed are contacted with the molecular sieve at a temperaturein the range of from 200° C. to 650° C. In a further preferredembodiment the temperature is in the range of from 250° C. to 600° C.,more preferably in the range of from 300° C. to 550° C., most preferablyin the range of from 450° C. to 550° C. Preferably the reaction toproduce the olefins is conducted at a temperature of more than 450° C.,preferably at a temperature of 460° C. or higher, more preferably at atemperature of 490° C. or higher. At higher temperatures a higheractivity and ethylene selectivity is observed. Molecular sieves havingone-dimensional 10-membered ring channels can be operated underoxygenate conversion conditions at such high temperatures withacceptable deactivation due to coking, contrary to molecular sieves withsmaller pores or channels, such as 8-membered ring channels.Temperatures referred to hereinabove represent reaction temperatures,and it will be understood that a reaction temperature can be an averageof temperatures of various feed streams and the catalyst in the reactionzone.

In addition to the oxygenate, and the olefinic co-feed, a diluent may befed into the reactor system. It is preferred to operate without adiluent, or with a minimum amount of diluent, such as less than 200 wt %of diluent based on the total amount of oxygenate feed, in particularless than 100 wt %, more in particular less than 20 wt %. Any diluentknown by the skilled person to be suitable for such purpose can be used.Such diluent can for example be a paraffinic compound or mixture ofcompounds. Preferably, however, the diluent is an inert gas. The diluentcan be argon, nitrogen, and/or steam. Of these, steam is the mostpreferred diluent. For example, the oxygenate feed and optionallyolefinic co-feed can be diluted with steam, for example in the range offrom 0.01 to 10 kg steam per kg oxygenate feed. In one embodiment smallamounts of water are added in order to improve the stability of thecatalyst by reducing coke formation.

In one embodiment, each gaseous effluent from one of the riser reactorstages, or preferably from all riser reactors individually, has aconcentration of C5+ olefins (pentenes and higher olefins) of below 10wt %, preferably below 5 wt %, more preferably below 2 wt %, yet morepreferably below 1 wt %, still more preferably below 0.1 wt %. Inparticular, the C5+ olefins can comprise at least 50 wt % pentenes, morein particular at least 80 wt %, even more in particular at least 90 wt %of pentenes. In particular the pentene concentration of the gaseouseffluent can be below 10 wt %, preferably below 5 wt %, more preferablybelow 2 wt %, yet more preferably below 1 wt %, still more preferablybelow 0.1 wt %.

In this way the ratio of C5+ olefins (in particular C5 olefins) tooxygenate at the subsequent riser inlet to which oxygenate is added iskept minimum in the process. Without wishing to be bound to a particularhypothesis, it is currently believed that keeping the ratio C5+olefins/oxygenate, in particular C5 olefins/oxygenate, small isbeneficial to ethylene selectivity, more in particular in the case thatthe oxygenate comprises oxygen-bonded methyl groups. It is currentlybelieved that pentenes should be preferentially cracked to yieldethylene and propylene, as opposed to alkylation to higher olefins byreaction with the oxygenate. Cracking of higher olefins is thought toresult is a lower concentration of ethylene in the final product.

The olefinic product or reaction effluent is typically fractionated. Theskilled artisan knows how to separate a mixture of hydrocarbons intovarious fractions, and how to work up fractions further for desiredproperties and composition for further use. The separations can becarried out by any method known to the skilled person in the art to besuitable for this purpose, for example by vapour-liquid separation (e.g.flashing), distillation, extraction, membrane separation or acombination of such methods. Preferably the separations are carried outby means of distillation. It is within the skill of the artisan todetermine the correct conditions in a fractionation column to arrive atsuch a separation. He may choose the correct conditions based on, interalia, fractionation temperature, pressure, trays, reflux and reboilerratios.

At least a light olefinic fraction comprising ethylene and a heavierolefinic fraction comprising C4 olefins and less than 10 wt % of C5+hydrocarbon species are normally obtained. Preferably also a water-richfraction is obtained. Also a lighter fraction comprising methane, carbonmonoxide, and/or carbon dioxide can be obtained, as well as one or moreheavy fractions comprising C5+ hydrocarbons. Such heavy fraction can forexample be used as gasoline blending component.

In the process also a significant amount of propylene is normallyproduced. The propylene can form part of the light olefinic fractioncomprising ethene, and which can suitably be further fractionated intovarious product components. Propylene can also form part of the heavierolefinic fraction comprising C4 olefins. The various fractions andstreams referred to herein, in particular the recycle stream, can beobtained by fractionating in various stages, and also by blendingstreams obtained during the fractionation. Typically, an ethylene and apropylene stream of predetermined purity such as pipeline grade, polymergrade, chemical grade or export quality will be obtained from theprocess, and also a stream rich in C4 comprising C4 olefins andoptionally C4 paraffins. In a preferred embodiment the process accordingto the invention is designed to produce lower olefins for recovery andonward processing and/or sale. Typically therefore, a stream comprisingat least 50 wt %, preferably at least 75 wt %, C2 to C3 olefins(ethylene and/or propylene) is separated from the reaction product,based on total reaction product.

It shall be clear that the heavier olefinic fraction comprising C4olefins, forming the recycle stream, can be composed from quantities ofvarious fractionation streams. So, for example, some amount of apropylene-rich stream can be blended into a C4 olefin-rich stream. In aparticular embodiment at least 90 wt % of the heavier olefinic fractioncomprising C4 olefins can be formed by the overhead stream from adebutaniser column receiving the bottom stream from a depropanizercolumn at their inlet, more in particular at least 99 wt % orsubstantially all.

Suitably the olefinic reaction effluent comprises less than 10 wt %,preferably less than 5 wt %, more preferably less than 1 wt %, of C6-C8aromatics. Producing low amounts of aromatics is desired since anyproduction of aromatics consumes oxygenate which is therefore notconverted to lower olefins.

The process may be started up by using olefins obtained from an externalsource for the olefinic co-feed, if used. Such olefins may for examplebe obtained from a steam cracker, a catalytic cracker, alkanedehydrogenation (e.g. propane or butane dehydrogenation). Further, sucholefins can be bought from the market.

When a molecular sieve having more-dimensional channels such as ZSM-5 ispresent in the oxygenate conversion catalyst, even in minority comparedto the molecular sieve having one-dimensional 10-membered ring channels,start up is possible without an olefinic co-feed from an externalsource. ZSM-5 for example is able to convert an oxygenate to anolefin-containing product, so that a recycle can be established.

Typically the oxygenate conversion catalyst deactivates in the course ofthe process. Conventional catalyst regeneration techniques can beemployed, such as oxidation of coke in a regenerator. The molecularsieve having one-dimensional 10-membered ring channels used in theprocess of the present invention can have any shape known to the skilledperson to be suitable for this purpose, for it can be present in theform of spray-dried particles, spheres, tablets, rings, extrudates, etc.Extruded catalysts can be applied in various shapes, such as, cylindersand trilobes. Spray-dried particles are preferred.

An embodiment of the present invention will now be described by way ofexample only and with reference to and as shown in FIG. 1, which is afront diagrammatic view of a catalyst return apparatus and riser reactorsystem in accordance with the present invention.

FIG. 1 shows a riser reactor system 10 comprising a first riser segment13, a second riser segment 14 and a third riser segment 15. The risersegments 13-15 are stacked such that they are co-linear so that fluidtogether with catalyst particles may travel up sequentially from thefirst riser 13 through the second riser 14 to the third riser 15. Theriser reactor system can be more than 40 meters high.

As described in more detail below, catalyst particles are separated fromthe gaseous reaction products in a separation zone 12 and recovered in acollection vessel or catalyst retention zone 16 at the top of the thirdriser segment 15. Catalyst C from the catalyst retention zone isreturned by the catalyst return apparatus 11 to the catalyst inlets atthe bottom of the three riser segments. The flowpath to the bottom ofthe first riser includes (in sequence) a conduit 29, a first stagevessel 28, a further conduit 27, a second stage vessel 17 and a finalconduit 23.

Valves 36, 31 are provided to control fluid flow in the conduits 29 and27, immediately upstream of the stage vessels 28, 17; and a third valve26 controls fluid flow in the conduit 23 immediately upstream of thefirst riser 13.

When fluidized catalyst flows down a standpipe for a significant length,the flow regime is difficult to control, in particular when a valve waspassed. The stage vessels allow the fluidized catalyst flow on the wayto the bottom of the riser to settle into a gas phase 17 a, 28 a, and afluidized solid phase 17 b, 28 b. From the fluidized solid phase in thestage vessel a well-defined flow into the subsequent section of theflowpath starts. In this way the stage vessels serve to reconstitutecatalyst flow.

The stage vessels 28 and 17 also have a much greater width than theconduits and are typically at least 2 m in width and at least 3 m inheight whereas the conduits can typically be 0.5 m-1 m in width. Thestage vessels suitably comprise gas inlets 34, 35 for the fluidisationof catalyst 11 during use. The stage vessels are also pressure balancedwith each other and the riser reactor system 10 via conduits 22 and 42.

Thus catalyst particles recovered from the top of the third riser 15 aretransported via the conduits 29, 27, 23 and stage vessels 17, 28 to thebottom of the first riser 13.

The presence of valves on the conduits and additionally also the stagevessels allows the pressure difference between the catalyst retentionzone 16 at the top of the riser reactor system, and the catalyst inletat the bottom of the riser reactor system 10, which can for example be 2bar or more, to be handled well, allowing reliable catalystrecirculation while observing the optimum operation window for thepressure difference over a valve of say maximum 1 bar.

The riser reactor system is especially useful for the catalyticconversion of oxygenates to olefins, especially C2 and C3 olefins.

Further features of the riser reactor system 10 and catalyst returnapparatus 11 will now be described. As shown in FIG. 1, a riser enddevice 20 is connected to a top end 21 of the third riser segment 15 toimprove separation of the catalyst from the reactor effluent from thethird riser segment 15, into the bottom of a collection vessel 16. A gasinlet 33 provides for fluidisation of the catalyst in the collectionvessel 16. Catalyst fines are recovered from the gaseous product by acyclonic separator 18 and the fines are returned to the collector vessel16. The product output from the cyclonic separator at line 19, such asincluding lower olefins, is recovered and processed further (not shown).A portion of the catalyst is diverted from the collection vessel 16 to aregenerator 39 via a conduit 40 and returned to the collection vessel 16via a conduit 43. A valve 41 on the conduit 40 provides for control ofthe catalyst sent to the regenerator 39 and a valve 42 on the conduit 43provides for control of the catalyst return from the regenerator 39.

Accumulated (fluidised) catalyst C is shown in the collector vessel 16,and also shown in the stage vessels 28, 17.

In this embodiment having three riser sections, fluidised catalyst isadded to the riser reactor system 10 not only at the bottom of the firstriser 13 but also at the bottom of the second 14 and third 15 riserreactors. Standpipe 37 provides fluid communication between thecollector vessel 16 and the bottom of the third riser segment 15 forthis purpose. This standpipe is less than 20 meters high, so that asimple design with a single valve 38 suffices. Standpipe 30 connects thefluidized bed 28 b in stage vessel 28 to the bottom of the second riser14. The combination of conduit 29, stage vessel 28 and conduit 32 withthe valves 32 and 36 represents an embodiment of a catalyst returnapparatus according to the invention in itself.

Valves 38 and 32 are provided on the conduits 37, 30 respectively tocontrol the flow of fluidised catalyst into the riser reactor system 10at these points.

Each of the conduits 23, 27, 29, 37 has a height of 20 m or less, inparticular a length of 20 m or less.

It will be understood that simpler embodiments of the present inventionare obtained in case the riser reactor system only contains a singletall riser instead of the three segments shown. At the same overallheight, the conduits 30 and 37 are not required then.

The system can also be simplified by not including the stage vessels 28and 17 as well as the lines 22 and 44 as well as 34, 35. If further onlya single tall riser was present, the catalyst return apparatus would beformed by the conduits 23, 27 and 29 with the valves 26, 31, 26, whichconduits would be in direct fluid communication with each other then.Depending on the total height, the one of the valves could even beomitted.

The present invention has been discussed in the context of anoxygenate-to-olefins reactor system. It will however be clear that theinvention can also be used in riser reactor systems for other processes,including other chemical processes, and refinery processes such asfluidized catalytic cracking.

What is claimed is:
 1. A process for reacting a feedstock in a riserreactor system comprising a riser reactor and a catalyst returnapparatus suitable for use in a riser reactor system, the catalystreturn apparatus comprising a catalyst return conduit and at least twoflow control devices in series along the catalyst return conduit and atleast two means to reconstitute catalyst flow, excluding regenerationand heat exchange units, in series along the catalyst return conduit,wherein the length of the catalyst return conduit is more than 20 m.,the process comprising: reacting a feedstock in the presence of acatalyst in the riser reactor to form a reaction product; separatingcatalyst from the reaction product and returning the catalyst to theriser reactor through the catalyst return apparatus comprising at leastone stage vessel; and holding a fluid comprising the catalyst in the atleast one stage vessel for a residence time of at least 10 seconds.
 2. Aprocess as claimed in claim 1, wherein the process comprises reacting anoxygenate feedstock in the presence of an oxygenate conversion catalystunder oxygenate-to-olefin conversion conditions in the riser reactor toobtain an olefinic product.
 3. A process as claimed in claim 1, whereinthe length of the catalyst return conduit is more than 30 m.
 4. Aprocess as claimed in claim 1, wherein the means to reconstitutecatalyst flow comprises at least one stage vessel, wherein the at leastone stage vessel has a width at least twice a width of the conduit.
 5. Aprocess as claimed in claim 1, wherein the means to reconstitutecatalyst flow comprises a gas inlet to further fluidise catalystparticles during use.
 6. A process as claimed in claim 1, wherein thecatalyst return conduit comprises at least one means to reconstitutecatalyst flow at least every 20 m.
 7. A process as claimed in claim 1,wherein at least one flow control device is provided for each means toreconstitute catalyst flow, within 2 m upstream thereof, to control thefluid within the conduit at that point.
 8. A process as claimed in claim1, wherein the conduit comprises a first portion above the means toreconstitute catalyst flow and a second portion below the means toreconstitute catalyst flow, wherein the first and second portions arenot co-linear.